This application is a continuation-in-part of application Ser. No. 250,660 filed Sept. 29, 1988 which teaches a two-stage process for upgrading olefinic light gas feedstock (termed "light gas" for brevity herein) containing C.sub.2 -C.sub.5 lower, particularly C.sub.3 -C.sub.5, olefins (alkenes) and paraffins (alkanes); and also, Ser. No. 184,465 filed Apr. 20, 1988 which teaches the oligomerization of C.sub.2.sup.+ olefins in a single reaction zone of catalyst, present as a turbulent fluid bed in which olefins in the super-dense phase are oligomerized (hence "super-dense" reactor).
This invention relates to a two-stage process utilizing a known oligomerization catalyst in each of two super-dense reactors in tandem, each of which operates to generate an effluent at sufficiently high temperature and pressure to be in the super-dense condition defined herebelow.
Because super-dense effluent from a MOG (for "Mobil Olefin to Gasoline") riser reactor (or, simply "riser") flows directly into the MODL (for "Mobil Olefin to Distillate and/or Lubes") reactor (referred to as such when reference is made to its operation either in a distillate mode to make a major fraction of distillate, or in a lubes mode to make a major fraction of lubes), the super-dense reactors truly operate in tandem; and, operation in tandem dispenses with a regenerator for spent catalyst (hence "regeneratorless", or the neologism "regenless" for ease, convenience and brevity). When the MODL secondary, fluid-bed reactor operates in a distillate mode, it is referred to herein as a "MOD" reactor; and, when this MODL reactor operates in a lubes mode, it is referred to as a "MOL" reactor. The MODL reactor may, under specific, generally atypical circumstances, also be operated to produce both gasoline and distillate in an effluent ("MODL effluent") which contains a larger proportion by weight (wt) of gasoline than is present when the reactor is operated in the distillate mode. Under such operating conditions the secondary reactor is referred to as a "MOGD" (for "Mobil Olefin to Gasoline & Distillate") reactor.
Operation of the regenless process with tandem super-dense reactors provides hitherto unattainable flexibility and stability of operation, particularly with respect to (i) the wide range of feedstocks which may be used, (ii) the efficiency with which a first, riser reactor may be operated at high severity in the transport mode, (iii) the economy inherent in the elimination of intercoolers and other equipment, including a regenerator, for "spent" catalyst, (iv) the ability to guard the catalyst in the MODL reactor against undue contamination which is countered in the MOG riser, and (v) the high conversion to distillate and/or lubes obtained at low severity in a fluid bed reactor operating under super-dense conditions in the turbulent sub-transport regime.
The extent to which the operation of a regenerator influences the complexity and cost of operating an olefin oligomerization process at a relatively much higher pressure than that of the regenerator, may be appreciated by studying the disclosure of our copending patent application Ser. No. 286,204 filed Dec. 19, 1988. In the particular instance where an oligomerization reactor is operated under super-dense conditions, both physical equilibria and practical considerations dictate that a spent catalyst stream be stripped, for example with steam, and regenerated at a relatively much lower pressure than that at which the super-dense fluid bed reactor operates. Such a process requires the use of lock-hoppers between the reactor and the regenerator, to lower the pressure, and, also between the regenerator and the reactor, to charge the regenerated catalyst from a low pressure zone to a much higher pressure zone.
Despite improvements in regeneration and the process scheme using lock-hoppers, regeneration is still a demanding and expensive unit operation, and much effort has been directed towards configuring a process which dispenses with the use of a regenerator, as for example, disclosed in our copending patent application Ser. No. 339,466, filed Apr. 17, 1989.
An ancillary consideration is the amount of carbonaceous residue ("coke") deposited on the catalyst as an undesirable byproduct of the oligomerization reaction The less deposited, the smaller the cost of regeneration. But the amount of coke deposited is a function of numerous operating conditions, and no prior art reference teaches how operation of a reactor in "plug" flow (characteristic of a riser reactor) might affect coke formation, as compared to coking up of the same catalyst in a turbulent fluid bed. Clearly, if no coke is formed, no regeneration would be required. But as long as there is substantial coke formation, operation of a continuous oligomerization process demands that some steps be taken to cope with the coke formation.
The first stage of the regenless process comprises upgrading either light gas, FCC gas, and/or, light naphtha, boiling range 175.degree. C. (347.degree. F.) to 240.degree. C. (464.degree. F.), any one of which contains at least 10 percent by weight (% by wt) olefins, to intermediate range hydrocarbons boiling in the range from 50.degree. C. to 204.degree. C. (125.degree. F.-400.degree. F.) ("gasoline") in a primary riser reaction zone.
The key to economic operation of our process is the excellent conversion, in excess of 90%, which we obtain despite operating a riser in the transport regime, because the fluidizing medium at least in the upper portion of the MOG riser reactor and throughout the fluid bed MODL, is neither gas nor liquid, but a supercritical fluid in the super-dense phase. The phase of the feed to the MOG riser is not critical and may be gas, or liquid under high pressure, the latter being immediately vaporized upon initiation of the exothermic oligomerization reaction. However, it is essential that the upper portion of the reactor be in the super-dense phase, and that the effluent leaves in the super-dense phase. Therefore the MOG riser will be referred to herein as being in the super-dense phase.
In contrast with a typical transport riser reactor operating in the gaseous phase, in which the suspended solids are in the range from 1 to about 5% by volume, we can operate a super-dense riser with solids in the range from 10 to 30%, and because of the physical properties of the super-dense fluidizing medium, we can maintain a satisfactory dispersion in each zone of the riser at much lower superficial velocities than in a prior art gas-phase transport regime.
We thus obtain the advantages of dilute-phase fluidization under transport conditions with much higher "solids density", that is, less voidage, than in a prior art transport zone. By "dilute-phase fluidization" I refer to a condition in which there is a net flow of fluid through the disperse suspension, but no net flow of solids; the particles move about in the suspension but do not flow along with the gas stream, or the reactor would empty. When the fluid velocity is further increased, the particles flow along with the gas at a particle velocity approximately equal to the differential increase in gas velocity. Under such transport conditions the reactor does empty; to maintian a solids inventory, particles must be continually fed to the reactor with the inlet fluid.
Though our '660 application disclosed that the MOG primary reactor used therein may be operated as a riser reactor, a fluid bed was used. Since the operating pressure was relatively low, operation of a riser at such pressure would not be expected to pose a problem. The concept of operating a riser, or any portion of a riser, under super-dense conditions simply did not occur to us because it was hard to conceive of plug flow under super-dense conditions. In plug flow, there is essentially no back-mixing in the axial direction, and very little, if any, mixing in the radial directions. The homogeneous distribution of catalyst, the isothermal conditions, and the narrow range of distribution of the hydrocarbon components in a turbulent fluid bed, are so different from the conditions in a riser reactor, that the added limitation of super-dense conditions made operation of the MOG riser unpredictable; and, tying its operation to a fluid bed MODL reactor, also operating under super-dense conditions, more so.
In particular, under plug flow conditions, there was no basis upon which we could predict how the catalyst would perform at the high WHSV required for a riser operating with a super-dense phase, necessarily with plug flow, and at pressure and temperature conditions at which it is critical that there be no liquid phase present. Nor did we fully appreciate the engineering requirements of controlling a riser operating in the super-dense phase, or, with the lower portion operating in the mixed phase and the upper portion operating in the super-dense phase.
The super-dense phase is defined by operating conditions such that no liquid may be present, or, above those at which liquid may be present, hence referred to herein as P.sub.max and T.sub.max. Such operating conditions prevail at near-critical and super-critical pressure and/or temperature in the super-dense phase which is always present in the upper portion of the riser, so that the effluent leaves under super-dense conditions. By "near-critical" we refer to a pressure which is typically at least 2857 kPa (400 psig), and a temperature which is typically at least 204.degree. C. (400.degree. F.); such conditions are always present for the effluent, but not necessarily always above the critical temperature of the feed. It will be appreciated that the pressure of the feed to the riser will always be above the operating pressure of the MODL if they are to operate in tandem. In other words, the reactor converts light gas to heavies in a single zone operating at a pressure and temperature outside a tightly circumscribed region of pressure and temperature ("critical P & T region") which region lies near, or above the apex of a phase diagram defining the critical point (P.sub.cr, T.sub.cr) of the mixture of hydrocarbons in the reactor.
Each super-dense reactor in this process operates best when the effluent from each is in the supercritical pressure and temperature region, at a pressure which is about 3550 kPa (500 psig) or above, and a temperature which is about 204.degree. C. (400.degree. F.) or above. The hope that the selectivity and yield of a riser under such conditions might be favorable, was tempered by the realization that a riser operating under high severity conditions would be deemed impractical from an economic point of view.
The super-dense MOG fluid-bed reactor in our '660 case, was operated at relatively low weight hourly space velocity ("WHSV", it being understood that WHSV signifies pounds of olefins fed per pound of zeolite per hour) but otherwise under process conditions generally within the ranges specified for those used in a process described in our U.S. Pat. No. 4,777,316, except that we operated the '660 MOG reactor to produce higher conversion to gasoline, and a "distillate-rich" gasoline (at least 1 part distillate for 10 parts by wt gasoline) effluent substantially free of aromatics (that is, less than about 3 mol percent aromatics), which effluent contained slightly more paraffins than in our '316 process. Further, the effluent in our '660 application was condensed and fractionated under conditions different from those in our '316 process so that we avoided sending (C.sub.10.sup.+ and heavier) components to the secondary reactor thus providing a tailored, olefin-rich C.sub.5.sup.+ feed, substantially free of distillate, to a secondary reactor in which the feed is converted either to distillate, or to lubes depending upon the particular preselected mode in which the secondary reactor is operated.
Though we also suggested using a super-dense MODL reactor we failed to recognize that, (i) directly flowing the "distillate-rich" effluent from the MOG reactor to the MODL reactor without changing the phase of the effluent, did not adversely affect selectivity and yield in the MODL reactor; and (ii) after high activity catalyst from the MODL reactor was "used" (or "spent"), the useful life of the "spent" catalyst as an oligomerization catalyst, could be beneficially, effectively prolonged, if this catalyst was directly transferred to the riser reactor.
Despite the overlap in the operating conditions of pressure and temperature for both the MOG and the MODL reactors in the disclosures of some of our preceding inventions, we did not recognize the "doability" of operating a super-dense MOG riser with plug flow (transport regime), let alone the benefit of doing so. Nor did we realize that we could do so without sacrificing the catalyst's selectivity and conversion of C.sub.2 -C.sub.5.sup.= when the riser was operated to produce a "distillate-rich" MOG effluent. Nor did we know the particular operating conditions for operating a fluid-bed MODL reactor with a tailored C.sub.5 -C.sub.9 feed which produces more distillate than gasoline.
"Distillate" refers to C.sub.10.sup.+ hydrocarbons boiling in the range from 130.degree. C. to 343.degree. C. (266.degree. F.-650.degree. F.); "lubes" refers to hydrocarbons boiling above 343.degree. C. (650.degree. F.) having a viscosity in the range from 4 cp to about 40 cp, measured at 100.degree. C. The particular operational mode chosen depends upon which particular boiling range of oligomerized product is desired, though in either the MOD or MOL modes, a minor amount of C.sub.5.sup.+ gasoline range hydrocarbons may also be formed. When this occurs, the gasoline, typically not a desired product in our process, is recycled to the MODL reactor to yield the desired distillate or lubes product. Light gas containing a substantial, preferably a major portion, typically more than 75% of combined propene and butenes, is a particularly well-suited feed to the reactor.
The specific embodiments of this invention derive from operating the MOG riser reactor as a recirculating ("recirc" for brevity) riser with partially "coked" catalyst which is obtained from the MODL fluid-bed reactor. Conditions in each reactor are such that only the olefins are oligomerized. Operation of the reactors in tandem provides the flexibility to operate the process to produce mainly distillate, or lubes, or even gasoline in the secondary reactor, but always producing a "distillate-rich effluent" from the MOG riser reactor. Tailoring the super-dense upper portion of the MOG riser to provide an effluent for the fluid bed MODL reactor results in a surprisingly effective combination of conventional unit operations which permit continuous operation of the process. In this "maximum conversion" operation of the MOG reactor, an exceptionally high conversion of light gas (or, FCC gas), or light naphtha to olefins is obtained at WHSV&gt;10 hr.sup.-1.
Feeding distillate-rich effluent formed in the MOG riser to the fluid bed MODL does not diminish the yield of distillate produced in the MODL despite some expected cracking of distillate in the MODL; and, after separating the distillate, the remaining gasoline-containing stream is recycled to the MODL reactor. But for this combination of a gasoline-containing recycle and distillate-rich MOG riser effluent to the MODL reactor, we would not have the unexpectedly economic oligomerization of olefins in the MODL reactor, along with beneficial processing flexibility and savings in the costs of operation, all of which help make the process economical.
Except for means to separate entrained catalyst in the distillate-rich effluent to the MODL reactor, we now dispense with equipment to process the MOG effluent to the MODL reactor. As explained in the '660 case, the thrust was to remove distillate from the distillate-rich effluent before flowing the gasoline to the MODL reactor. A process scheme to do so required providing a debutanizer, and placing a gasoline/distillate splitter ("G/D splitter") or a high temperature separator ("HTS") before the MOD reactor in the distillate mode. We dispense with the equipment.
Developments in fluid-bed and fixed bed catalytic processes using a wide variety of zeolite catalysts have spurred interest in commercializing the conversion of olefinic feedstocks to C.sub.5.sup.+ hydrocarbons including gasoline, diesel fuel, lubes, etc. In addition to the discovery that the intrinsic oligomerization reactions are promoted by aluminum metallosilicate (hereafter, "ZSM-5 type") zeolite catalysts, several discoveries relating to implementing the reactions in an apt reactor environment, have contributed to the success of current processes. These are environmentally acceptable processes for utilizing feedstocks containing lower olefins, especially C.sub.3 -C.sub.5.sup.= (olefins), though a significant quantity, up to 40% ethylene, along with olefins and paraffins heavier than C.sub.5 may also be present. A predominantly olefinic light gas containing more than 50% by wt, and preferably more than 60%, of combined propene and butenes, is a particularly well-suited feed to oligomerization reactors using a ZSM-5 type catalyst. It will be recognized that the higher the content of C.sub.2 H.sub.4 and C.sub.3 H.sub.6 in the feed to the MOG riser, the higher its operating pressure.
In our MOG+MODL combination of tandem super-dense riser+fluid-bed reactors, it is essential that the former operates with a relatively lower activity catalyst than the latter, at WHSV&gt;10 hr.sup.-1, under relatively high (top) temperature conditions in its upper half; and that the latter operates with a relatively higher activity catalyst than the former, at WHSV&lt;10 hr.sup.-1, under relatively low temperature conditions.
In addition to the operational flexibility referred to hereinabove, afforded by the combination of tandem super-dense riser+fluid-bed reactors, our process results in a sufficiently low "coke-make" to permit regenless operation. A slipstream of spent catalyst is either intermittently or continuously withdrawn from the riser reactor and flowed to a fluid catalytic cracking (FCC) cracker.
Dispensing with spent catalyst from the MOG riser in this fashion is practical because of the favorably low coke make, and the high conversion which results from operation above P.sub.max and at or above T.sub.max ; also, because the entire contents of the MOG riser is in the transport regime, and the fluid-bed is in a turbulent regime, the solid acts both as catalyst and heat transfer medium to maintain isothermal conditions. In this process, the superdense fluid is neither gas nor liquid, but for convenience and familiarity, we treat the oligomerization reaction as being a gas/gas reaction.
More particularly, the MOG riser reactor operates continuously to oligomerize light gas containing propene, butenes and pentenes, preferably in the absence of added hydrogen, to a C.sub.10.sup.+ rich hydrocarbon stream, with higher pressure in the riser than the MODL fluid-bed, whichever its mode of operation.
In the MOD mode, the reactors are operated at relatively low pressure in the range from about 2857 kPa to about 10436 kPa (400 psig-1500 psig), and relatively high temperature in the range from 260.degree. C. to about 371.degree. C. (500.degree. F.-700.degree. F.). In the MOL mode the reactors are operated at relatively high pressure in the range from about 5270 kPa to about 13881 kPa (500 psig-2000 psig), and relatively low temperature in the range from 204.degree. C. to about 315.degree. C. (400.degree. F.-600.degree. F.). Even higher pressures, as high as 20821 kPa (3000 psig) may be used if the economics of operating at such high pressure can be justified by the lube "make".
The combination of MOG and MODL reactors in tandem is uniquely effective because the MOG reactor functions as a "guard" reactor for the MODL. Because of the sensitivity of a ZSM-5 type of catalyst to basic nitrogen-containing organic compounds such as alkylamines (e.g. diethylamine), or, to oxygenated compounds such as ketones, it is important to protect the catalyst in the MODL reacator. It will be recognized that alkylamines are used in treating light gas streams, and ketones are typically present in Fischer Tropsche-derived light ends streams, both of which streams are particularly well-suited for upgrading by oligomerization. This sensitivity (poisoning), is a characteristic of the catalyst under the process conditions of prior art olefin oligomerization processes, particularly the fixed bed processes operated at high pressure. Such processes require the addition of hydrogen as a preventitive antidote. Though our process is not adversely affected by the presence of hydrogen, there is no readily discernible economic incentive for using hydrogen in either the primary-stage or secondary-stage reactors, and we prefer not to do so.
Though the earliest prior art, moderate-pressure processes, for example those disclosed in U.S. Pat. Nos. 3,827,968 and 3,960,978 to Givens et al, used a zeolite catalyst to oligomerize lower olefins under moderate conditions, and produced excellent conversions to distillate range olefins in a fixed bed microreactor, some over-riding problems relating to operating the process economically were not foreseen (see "Conversion of C.sub.2 -C.sub.10 Olefins to Higher Olefins Over Synthetic Zeolite ZSM-5" by W. E. Garwood presented at the Symposium on Advances in Zeolite Chemistry before the Division of Petroleum Chemistry, Inc., American Chemical Society, Las Vegas Meeting Mar. 28-Apr. 2, 1962).
The '978 patent discloses that low alpha ZSM-5 and ZSM-11 catalysts not only have reduced activity for cracking n-hexane and other paraffins, but also produce less than 10% by wt aromatics. The runs were made in a fixed bed microreactor, and, at that time, it was not known that the process required the addition of hydrogen to control coke deposition and to prevent poisoning of the catalyst by nitrogen-containing organic impurities. The basic knowledge that low activity ZSM-5 and ZSM-11 type catalysts effectively oligomerized lower olefins, was used to arrive at improvements in "Catalytic Conversion of Olefins to Higher Hydrocarbons" in U.S. Pat. No. 4,456,779 to Owen et al. which discloses oligomerization of olefins in a MOD reactor system of three downflow fixed beds, in series, with intercoolers; and, more recently, in "Conversion of LPG Hydrocarbons to Distillate Fuels or Lubes Using Integration of LPG Dehydrogenation and MOGDL" in U.S. Pat. No. 4,542,247 to Chang et al which discloses fixed beds in a two-stage catalytic process for converting paraffins to olefins which in turn are converted to gasoline and distillate. The first stage MOG reactor is operated under conditions given in U.S. Pat. Nos. 3,960,978 and 4,211,640 to Givens et al. Under these conditions there is a substantial make of aromatics which are undesirable if the effluent from the MOG is to be converted to distillate (aromatics lower the cetane number, among other things).
In the '779 process, multiple fixed bed reactors are used, each operating in the same range of process conditions, and it was essential to dilute the feed to the reactors with both lower alkanes and recycled gasoline, to maintain a controllable exotherm in the bed. To provide the gasoline recycle, the effluent from the operating reactors (a spare reactor is always being regenerated) is debutanized after oligomerization of olefins is completed. Moreover, the fixed-bed processes in both the '247 and '779 patents require the addition of hydrogen for the reasons given hereinabove. Thus, despite operation at as high a pressure as is economically feasible, the use of hydrogen with a high concentration of lower alkanes dictates that the oligomerization be carried out in the gaseous phase, or vapor/liquid phases, thus aggravating both the heat transfer and mass transfer problems. When we use a fluid-bed MODL reactor, it operates with the hydrocarbons in the super-dense phase, the precise conditions of operation, being determined by economics.
Because Chang et al first dehydrogenated a paraffinic feed, they typically converted 30-40% of the paraffins to olefins. The feed to the MOG reactor therefore was predominantly C.sub.3 /C.sub.4 paraffinic, as was the effluent from the MOG reactor, since the undehydrogenated C.sub.3 /C.sub.4 paraffins are not oligomerized. Because, after oligomerization in the '247 fixed bed MOG reactor, the effluent still contained a major proportion of C.sub.3 /C.sub.4 paraffins, Chang et al had to separate the paraffins from the olefins in the effluent (so that the separated C.sub.4.sup.- paraffins could be recycled to be dehydrogenated). Since, under their conditions, the make of C.sub.10.sup.+ components was relatively small, they failed to realize the criticality of separating the C.sub.10.sup.+ components before the effluent from the MOG reactor was further oligomerized.
Though neither Owen et al, nor Chang et al, knew it at the time, in practice, a fixed bed requires the addition of a substantial quantity of hydrogen (for the reasons given), which fixed bed nevertheless is far less effective than a fluid bed for the specific purpose of "cleaning up" the MOG effluent. It is this volume of hydrogen which adds to the already large volume of diluents being used as a heat sink, albeit an inefficient one. Nothing in either the '779 or the '247 patents suggests the surprising benefits of operating with a fluid bed in the absence of added hydrogen and fluidized with a feed containing too little alkanes to serve as a significant heat sink, namely less than about 50% by wt, preferably less than 30% by wt.
The earlier references disclosed that the product distribution from an MOGD reactor may be tailored by controlling process conditions, such as temperature, pressure and space velocity. Gasoline (C.sub.5 -C.sub.10) is readily formed at elevated temperature (preferably about 400.degree. C.) and pressure from ambient to about 2900 kPa (420 psia), preferably about 250 to 1450 kPa (36 to 210 psia). Olefinic gasoline could be produced in good yield and may be recovered as a product; or, it could be fed to a low severity, high pressure reactor system for further conversion to heavier distillate-range products. Distillate mode operation could be employed to maximize production of C.sub.9.sup.+ aliphatics by reacting the lower and intermediate olefins at high pressure and moderate temperature. Operating details for typical MOGD oligomerization units are disclosed in U.S. Pat. Nos. 4,456,779 and 4,497,968 (Owen et al); 4,433,185 (Tabak); 4,456,781 to Marsh et al; and U.S. Pat. No. 4,746,316 to Avidan et al.
None of the foregoing alternatives disclosed the technical and economic difficulties of operating a riser MOG reactor under super-dense conditions, or of the fluid-bed MODL reactor also under super-dense conditions, in tandem with the riser MOG reactor.
The combination of tandem super-dense MOG riser+MODL fluid-bed reactors is unexpectedly effective because the riser rids the feed of poisons just as effectively as if it was a fluid-bed, while operating under high severity conditions which nevertheless produce the maximum conversion of olefins to C.sub.5.sup.+ olefins, substantially free of aromatics, economically. Since, in addition to the MOG riser effluent, only a gasoline-containing stream (after separation of the distillate) is to be recycled to the fluid bed MODL reactor, poisoning of the MODL catalyst is essentially negated. Because a substantial portion of the coke formation takes place in the MOG fluid bed, our MODL reactor operates with so little coke deposition that spent MODL catalyst can be reused in the MOG riser. Regeneration of the spent catalyst from the riser reactor can be avoided by reusing the spent MOG catalyst in the FCC reactor.
U.S. Pat. Nos. 4,417,086 and 4,417,087 to Miller teach a two-zone reactor operating in the transport mode where the relative superficial gas velocity is greater than the terminal velocity in free fall. Though the operation of a fluid-bed is illustrated (example 2 in each of the '086 and '087 patents) note that no operating pressure is stated in the former, and that operating pressure in the latter is 10 psig (24.7 psia, 170 kPa). The general disclosure that the processes may be operated at a pressure in the range from subatmospheric to several hundred atmospheres, but preferably 10 bar or less, and most preferably 0 to 6 bar, (see middle of col 6 in '086, and, near top of col 5 in '087) is not so ingenuous as to be meant to apply equally to the fixed bed (example 1 of '086 and '087, each illustrates 34.5 bar, 500 psi) and the 170 kPa fluid-bed.
In U.S. Pat. Nos. 3,960,978 and 4,021,502, Plank, Rosinski and Givens disclose conversion of C.sub.2 -C.sub.5 olefins, alone or in admixture with paraffinic components, into higher hydrocarbons over crystalline zeolites having controlled acidity. Garwood et al have also contributed to the understanding of catalytic olefin upgrading techniques and have contributed improved processes as in U.S. Pat. Nos. 4,150,062, 4,211,640 and 4,227,992. The '062 patent discloses conversion of olefins to gasoline or distillate in the range from 190.degree.-315.degree. C. and 42-70 atm; and this, and the '640 and '992 disclosures are incorporated by reference thereto as if fully set forth herein.